Catalytic conversion process for the production of low freezing point fuels



United States Patent 3,110,662 CATALYTEC CtENVERSlGN PROCESS FOR THE PRQDUCTESN 0F LQW FREEZING PGINT FUELS John W. Scott, .Ira, Ross, and Robert H. Kozlowski,

Berkeley, Caiitl, assignors to California Research Corporation, San Francisco, Caiifi, a corporation of Bellaware No Drawing. Filed June 2.4, 1960, er. No. 38,458

1 Qiairn. (ct. 2ss ss) This invention relates to a hydrocarbon conversion process and, more particularly, to a process for the catalytic conversion of petroleum distillates to produce fuels characterized by low, freezing points. I

More specifically, this invention is directed to a hydrocarbon conversion process whereby a first stage pretreating step is employed to prepare an essentially nitrogen-free hydrocarbon feed which in turn is passed into a second stage isocracking step wherein the feed converted to fuels having low freezing points which are suitable for use as jet, diesel and/ or other fuel fractions. The second stage step is conducted in the presence of hydrogen and a selective catalyst composition incorporating particular hydroge-nating components disposed on a highly active, solid, acidic catalytic support at elevated pressures and relatively low temperatures, resulting in a substantial consumption of hydrogen. This second stage operation can be considered generally to involve low temperature selective hydrocracking. For convenience and descriptiveness, this hydrocracking reaction is herein re ferred to as isocracking.

The present process, directed to the production of a low freezing point fuel, comprises pretreating a nitrogenous hydrocarbon distillate feedstock having an ASTM 13-1160 5% point above about 475 F. and boiling over a range of at least 50 F., in a pretreating zone to reduce its total nitrogen content to below about p.p.m. A substantial proportion of the resulting treated feedstock, along with at least 1500 s.c.f. of hydrogen per barrel of the treated feedstock, is contacted in an isocracking zone with a catalyst comprising a hydrogenating component selected from the group consisting of compounds of nickel and cobalt disposed on an active, solid, acidic catalyst support at a hydrogen partial pressure of at least 350 p.s.i.a., at a temperature between 350 and 700 F. and at a liquid hourly space velocity of above 0.2 v./v./hr., the reaction in said isocracking zone being characterized by a net consumption of hydrogen. The portion of the effluent from said isocracking zone boiling over the range of from 360 F. to the lowest of the two temperatures represented by the ASTM D1160 5% point of the untreated feedstock and 525 F., being a fuel characterized by an initial freezing point of less than about 40' F. and comprising at least 10 volume percent of the total C-,+ synthetic portion of the efiiuent of said isocracking zone.

As used herein, the term ASTM D-1160 5% point refers to the temperature at which 5 volume percent of the change is distilled overhead when the D1160 distillation procedure outlined in the volume ASTM Standards on Petroleum Products, American Society for Testing Materials, December 1958, pp. 614-627, is employed to determine the boiling range of the feed charge. Inasmuch as the procedure is used for high boiling feeds, thereby leading to minor inaccuracies in the measurement of the lower boiling end of the charge, the ASTM Dl 160 5% point is the determined temperatures -1() F. Thus, in reporting a 5% point as, say, 500 F., this point can vary from 490 to 510 F. Additionally, the term initial freezing point refers to the temperature at which the first crystal is formed, as determined by the Institute of Petroleurn method 16B.

The feedstocks employed in the subject process are 3,110,662 Patented Nov. 12, 1963 those nitrogenous hydrocarbon fractions having the specific distillation points and boiling range spread stated above.

. characteristics are those distillates normally defined as heavy straight-run gas oils and heavy cracked cycle oil. The feeds may be of straight-run origin as obtained from petroleum, or they may be derived from various processing operations and, in particular, from thermal or catalytic cracking of stocks obtained from petroleum, gilsonite, shale, coal tar, or other similar sources.

One of the important variables in the conduct of the present process which has a material effect and permits the production of the desired products is the control of the nitrogen content of the charge stock to the isocracking stage. As indicated, the nitrogen level, expressed as total nitrogen, should be below about 10 p.p.m., although appreciably improved results are obtained if the nitrogen content is reduced to levels below 2 ppm, and this latter figure is preferred as the upper limit on nitrogen content of the feed to the isocracking zone. These low nitrogen levels are reached by subjecting the nitrogenous feedstock to denitrification by a pretreating step.

In general, the effect of a total nitrogen content in excess of 10 ppm. in the isocracking step is a reduction in catalyst activity which is reflected in reduced operational efliciency and poorer product distribution. As the nitrogen content increases above the specified maximum, higher reaction temperatures are necessary to maintain an economic per-pass conversion level. These higher reaction temperatures cause a disproportionate increase in the amount of product converted to gases and carbonaceous residues deposited on the catalyst surface and thus further decrease catalyst activity. Such further decrease in catalyst activity must be compensated for by resort to still higher operating temperatures if acceptable conversion is to be maintained and, thus, the o-n-stream life of the catalyst is shortened as limiting temperatures of about 700 F. are reached much sooner than would otherwise be the case.

The marked effect of nitrogen upon the isocracking reaction is in sharp contrast to that observed in conventional hydrocracking operations, which are conducted at higher temperatures. In such operations, the effect of nitrogen in the feed, even when present in substantial amounts, is small at reaction temperatures of about 800 F. and becomes scarcely noticeable at temperatures above 850 F. Under the low temperature conditions requisite to the present isocracking operation, on the other hand, nitrogen compounds present in the feed drastically reduce the etfectiveness of the catalyst.

As noted above, the feedstock is first subjected to a pretreating operation that is relatively selective for the removal of nitrogen compounds. The requisite low nitrogen levels may be reached, for example, by intimately contacting the feedstocks with various acidic media such as liquid acids (H etc.) or, in the case of feeds that are comparatively low in nitrogen compounds, with such solid acidic materials as acid ion exchange resins and the like. However, it is preferred to carry out denitrification by catalytic hydrogenation (hydrofining) of the feed. This entails contacting the feed at temperatures of from about 400 to 900 F. (preferably from about 500 to 800 F.), pressures of at least 300 p.s.i.g., liquid hourly space velocities (-LHSV) of from about 0.3 to 5 v./v./hr., along with atleast 500 s.c.f. of hydrogen per barrel of feed, with a sulfur-resistant hydrogenation catalyst. Any

of the known sulfactive hydrogenation catalysts may be used in the hydrofining pretreatment. The preferred catalysts have as their main active ingredient one or more oxides or sulfides of the transition metals, such as cobait, molybdenum, nickel and tungsten. These various materials may be used in a variety of combinations with In general, suitable feedstocks that possess such.

or without such stabilizers and promoters as the oxides and carbonates of K, Ag, Be, Mg, Ca, Sr, Ba, Ce, Bi, Cr, Th, Si, Al and Zr. These various catalysts may be employed per se or in combination with various conventional supporting materials. Examples of the latter are charcoal, fullers earth, kieselguhr, silica gel, alumina, bauxite, and magnesia. While any of the noted classes of conventional sulfactive hydrogenation catalysts may be employed, it has been found that a molybdenum oxide catalyst promoted by a minor amount of cobalt oxide and supported upon an activated alumina or a tungsten sulfide on activated alumina are suitable catalysts for the hydrofining operation. Another desirable catalyst is composed of molybdenum sulfide promoted by a minor amount of nickel sulfide supported on activated alumina. The catalyst may be in the form of fragments or formed pieces such as pellets and cast pieces of any suitable form or shape.

Following the pretreating operation, the efiiuent from this zone is treated so as to remove acid sludge (in the case of acid pretreatment) or, in the case where hydrofining is employed as the pretreating step, ammonia, therefrom. When hydrofining for nitrogen reduction, the preferred ammonia removal method involves injecting water into the total efliuent from the hydrofining reactor and passing the resulting mixture into a high pressure separator operating under such conditions of temperature and pressure (for example, 100 F. and 950 p.s.i.g.) that a gaseous overhead is removed that is predominantly hydrogen but which also normally contains some hydrogen sulfide and light hydrocarbons. This overhead can be recycled to the hydrofining reactor. Within the separator are formed two phases, an upper hydrocarbon phase and a lower aqueous phase. The latter, containing essentially all of the ammonia present (and some of the H 8 if present) in the effluent, is removed from the system. The hydrocarbon layer is then passed into a stripper or into a distillation column wherein any traces of hydrogen sulfide, ammonia and water are removed overhead.

A substantial proportion (preferably all) of the treated feedstock of the normally liquid effluent from the pretreating zone is passed into an isocracking zone wherein it is contacted, along with added hydrogen, with a catalyst comprising compounds of nickel and/or cobalt as a hydrogenating component intimately associated with a solid, active, acidic catalyst support under elevated temperatures and pressures. The catalyst support may comprise any one or more of such acidic materials as the conventional cracking catalysts containing composites of silica-alumina, silica-magnesia, silica-alumina-zirconia, acid-treated clays, and the like. In addition, satisfactory results can be obtained with, for example, acid-activated aluminas and synthetic metal aluminum silicates (such as the synthetic chabazites commonly referred to as molecular sieves) that impart the necessary cracking activity to the catalyst. Preferred cracking catalysts employed as supports are synthetically prepared silica-aluminas having silica contents in the range of from about 40 to 99 percent.

As noted, the hydrogenating component of the catalyst may comprise one or more of the compounds of nickel and/ or cobalt such as the sulfides, oxides and complexes of various metals such as cobalt-chromium and nickelchromium which can be considered, for purposes of description, as compounds of nickel and cobalt. Oddly, it has been found that catalysts composed of metallic nickel and/ or cobalt as the hydrogenating component are considerably less active than the compounds of such metals and are, accordingly, outside of the scope of the present invention. Of the various nickel and cobalt compounds, it has been found that, in general, the sulfides are the most active and, for that reason, are preferred.

The total amount of the hydrogenating component associated with the support may be varied within relatively wide limits of from about 0.1 to 35 percent (as d the metal), based on the weight of the entire catalyst composition.

The catalysts preferred for use in the isocrackrng step of the process of the invention are characterized by high activity; that is, they are capable of converting substantial proportions of the hydrocarbon feeds, above described, to lower boiling products under mild operating conditions. For test purposes only, the catalysts can be characterized by their ability to convert a particular light catalytic cycle oil to lower boiling products. Although the test feed does not fall within the scope of the present process, such readily available feeds can be easily employed to test catalyst activities. A catalyst having a high activity to convert the test light cycle oil feedstock will also have a high activity for the conversion of the heavier feedstocks employed in the process of the present invention. The preferredcatalysts are those that have the ability to convert 50 volume percent of a hydrofined typical light catalytic cycle stock containing less than '10 volume percent of components boiling below 400 R, in a true boiling point distillation, and having the following inspections.

ASTM distillation D-1160:

10% point, F 425 to 450 point, F 500 to 560 Aniline point, F 85 to 120 Basic nitrogen content, ppm BelowS to products boiling below about 400 F. at temperatures below 650 F., a hydrogen partial pressure of 1100 p.is.i.a., and a liquid hourly space velocity of 2.0 v./v./ hr., hydrogen being introduced into the reaction zone at the rate of 6500 standard cubic feet of hydrogen per barrel of feed. Catalysts exhibiting this degree of activity can be prepared in various ways. For example, catalysts of such activity can be prepared by impregnating a synthetic silica-alumina cracking catalyst support with sufficient nickel nitrate to give the impregnated silica-alumiha a nickel content in the range of about 6 to 15% by weight. The nitrate is decomposed and the impregnated support is then sulfided by contacting it with hydrogen sulfide or with hydrogen and a low molecular weight mercaptan or organic sulfide, or with hydrogen and a hydrocarbon containing a dissolved sulfur compound, at temperatures below about 750 F., and preferably below 700 F. Catalysts exhibiting the preferred activity can also be prepared by impregnating a silica-alumina support with a nickel compound, drying the impregnated support, and then heating (thermactivating) it in air to a temperature in the range l200 to 1600 F. for a period of about .25 to 48 hours. After the treatment, the catalyst is preferably sulfided in the manner indicated above at temperatures below 750 F. The preferred activity level can also be reached by treating cobalt sulfide or nickel sulfide supported on silica-alumina with hydrogen fluoride to impart of the order of at least 0.5 weight percent of fluoride in the catalyst.

In the operation of the isocracking step, the feedstock can be introduced to the reaction zone as either a liquid, vapor, or mixed liquid-vapor phase, depending upon. the temperature, pressure, proportions of hydrogen and boiling range of the charge stocks utilized. The feedstock is introduced in admixture with at least 1500 s.c.f. of hydrogen per barrel of total feed (including both fresh as well as recycle feed). Generally, at least 500, and normally from about 700 to 2000, s.c.f. of hydrogen are consumed in the isocracking reaction zone per barrel of total feed converted to synthetic products, i.e., products whose point is below the 5% point of the feed to the isoeracking zone under the same conditions of distillation. The hydrogen stream admixed with incoming feed is conventionally made up of recycle gas recovered from the efiluent from isocracking zone, together with fresh make-up hydrogen. The hydrogen content of the 3,11o,sea

recycle gas. stream in practice generally r'anges upward of 70 volume percent.

The pressures employed in the isocracking zone are in excess of a hydrogen partial pressure of at least 350 p.s.i.a., and may range upwardly to as high as 2000 p.s.i.a., with a preferred range being a hydrogen partial pressure of from about 500 to 1500 p.s.i.a. Hydrogen partial pressures below about 350 p.s.i.a. result in decreased conversions per pass, as well as shorter on-st-ream e-riods by reason of an accelerated tendency to induce dehydrogenation of naphthenes to aromatics, a reaction accompanied by carbonaceous fouling of the catalyst.

Generally, the isocracking zone feed may be introduced into the reaction zone at a liquid hourly space velocity (LHSV) of from about 0.2 to 5 volumes of hydrocarbon (calculated as liquid) per superficial volume of catalyst (v./v./hr.), with -a preferred rate being. from about 0.5 to 3 LHSV.

One of the most advantageous aspects of the subject isocracking process is that the reaction temperature is maintained below 700 F. The importance of such low temperature operations is reflected in long on-stream periods extending over many hundreds of hours and the production of extremely low yields of undesirable light gases.

In the preferred practice of this invention, the tempera ture at which the isooracking reaction is initiated when placing a cfiresh charge of catalyst (in-stream should be as low as possible (commensurate with the maintenance of adequate per-pass conversion levels), since the lower the starting temperature, the longer will be the duration of the said oil-stream period. For any given conversion, the permissible starting temperature is a function of catalyst activity since the more active catalyst (i.'e., those capable of effecting a relatively high per-pass conversion under given operating conditions such as the activated catalysts discussed hereinbefore) permit the unit to be placed onstream at lower starting temperatures than would otherwise be the case. In any event, the conversion reaction should be conducted at temperatures below about 700 F., with preferred initiating temperatures being in the range of from about 400 to 650 F. or even lower.

The isocracking stage of the present invention is conducted under the specified conditions of temperature, pressure and space rate such that at least 30 volume percent of the initial feed to the isocracking zone is converted per pass to synthetic products. Preferably, the reaction conditions are adjusted such that the per-pass conversion to synthetic product is in the range of from about 40 to 90 volume percent. Additionally, it is preferred to operate the isocracking process by periodically increasing the reaction temperature so as to maintain the selected per-pass conversion at relatively constant levels.

The isocracking stage of the present invention is well adapted to be carried out using any type feedcatalyst contacting method. Thus, such methods as fixedbed, moving-bed, slurry, or fluid catalyst systems can be employed by procedures well known in the art. The preferred method is that employing at least one fixed catalyst bed. Since reaction on-stream periods are so long, generally it is more economic to merely replace the deactivated catalyst with fresh catalyst. However, catalyst regeneration can be performed, for example, by contacting the deactivated nickel and/or cobalt compound containing catalyst with an oxygen-containing gas at temperatures of from about 700 to 1000 F. In the case of a sulfided catalyst, regeneration can be done by carrying out the noted burning operation, reducing the resulting nickel and/ or cobalt oxide to the metal and then sulfiding (in situ if desired) by contacting the catalyst at temperatures below 750 F. with hydrogen and H 8 or a gaseous compound capable of generating H 8. In the sulfide case, it may be desirable to eliminate the reduction step by sulfiding the oxide directly.

Because of the conversion in the isocracking zone of 6 atleast 30 volume percent of the feed to that zone, the total efiluent from the isocracker will contain products (exclusive of hydrogen) boiling over the entire range from methane up to about the 95% point of the feed to the unit. However, it has been found that, irrespective of any particular separation of the isocracker effiuent, if the (feedstock and reaction limitations of the present invention are observed, that the portion of the total isocracking zone eflluent boiling over the range of 360 to 525 F. or to the ASTM D-1160 5% point of the initial untreated feed, whichever is lowest, will be a fuel characterized by an initial freezing point below 40 F.

This result is believed due to the selectivity of the isocracking catalyst and reaction conditions to hydrocrack the other common molecular species much more readily than normal paraffins, which have high freezing points. Thus, it has been found that the isocracking reaction does not crack normal parafiins at near the rate it does other species, and further, that virtually no heavy normal paraflins are produced in the isocracker with the result that essentially no normal parafiins appear in the synthetic portion of the isocra-cking zone effluent boiling above 360 F.

In the following examples, the results of conducting the subject process with three different feedstocks under various modes of operation are shown. The properties of the raw, untreated feedstocks employed in these runs are shown in Table I below. Feedstock A was a heavy recycle oil derived from a fluid catalytic cracking unit, fed a mixture of Southern California, Four Corners and Minas gas oils. Feedstock B was a heavy cycle stock from a catalytic cracking process fed a mixture of Texas and Arabian gas oils. Feedstock C was a heavy straightrun Arabian gas oil.

TABLE I Feedstock A B C Gravity, API 24. 6 22. 6 29. 1 Aniline Point, F 137 161 177 Distillation, ASTM D-1160, F.

Start/5% 380/510 565/631 541/643 550/583 650/680 659/690 620 704 702 654/710 732/772 722/760 740/783 790/840 776/805 6, 200 16, 3, 300 900 566 462 Pour Point, F +40 +70 +60 Example 1 Feedstock A was reduced in nitrogen content (pretreated) in a multistage hydrofining operation by contact, along with 6,000 s.c.f. of hydrogen per barrel of feedstock, with a hydrofining catalyst composed of 19 weight percent (as the metal) molybdenum oxide and 5.9 weight percent (as the metal) cobalt oxide on an alumina support at a pressure of 1200 p.s.i.g. The first stage was operated at a maximum catalyst temperature of 745 and an LHSV of 0.5 v./v./ hr. The efiiuent from the first hydrofining stage was then repassed at a catalyst temperature of 700 F. and an LHSV of 0:8 v./v./hr. The properties of the nitrogen reduced, normally liquid efiiuent from the hydrofining zone are given in Table II below.

The normally liquid eifiuent from the hydrofiner and a recycle stream boiling'above about 500 F, along with 6500 set. (standard cubic feet) of recycle hydrogen per barrel of fresh feed plus recycle feed, were continuously passed into an isocracking zone and contacted with a thermactivated catalyst (120 ml.) composed of nickel sulfide (6 weight percent as nickel) disposed on a crushed (8 to 14 mesh) synthetic cracking catalyst support containing about 87 weight percent silica and about 13 weight percent alumina. The contacting operation was conducted at a liquid hourly space velocity (LI-ISV) of 0.8 v./v./hr., a total reaction pressure of 1200 psig. (hydrogen partial pressure of about 1100 p.s.i.a.) and a temperature of from 546 to 556 F. for a period over 100 hours. Sixty volume percent of the total feed to the isocracking zone were converted per-pass to products boiling below about 500 F. The efiluent from the isocracking zone was continuously recovered and all of the efiluent boiling above about 500 F. was recycled to the isocracker, as was the recovered hydrogen. Hydrogen consumption in the isocracking zone amounted to 1208 set. of hydrogen per barrel of total feed converted to products boiling below 500 F. The yields of the hydrocarbon efiluent boiling below 5 F. are given in Table Ill below.

The portion of the isocracking zone efiluent boiling over the entire range of from 360 to 500 F., amounting to 33.2 volume percent of the fresh feed converted, 31.5 volume percent of the (3 synthetic portion of the efiluent, and over 20 volume percent of the entire isocracker efiluent, had an initial freezing point of less than -72 F. In this particular operation, the lowest temperature represented by the temperature 525 F. and the ASTM D-l160 5% point of the initial untreated feed was the latter, said 5% point being about 510 F.

Example 2 In this example, inserted for comparative purposes, if the same feedstock and conditions as in Example 1 were employed except that 60 volume percent of the total feed to the isocracking zone were converted per-pass to products boiling below 525 F., rather than below 500 F., as in Example 1, the portion of the isocracldng zone efiluent boiling over the entire rmige of from 360 to 525 F, amounting to about 25 volume percent of the entire isocracker efiluent, would be expected to have an initial freezing point of only 30 F.

From the above two examples, it can be seen that, if a fuel boiling above 360 F. and having an initial freezing point below 40 F. is to be produced from this particular feedstock and recovered from the effluent of the isocracking zone, this fuel must not boil appreciably above the temperature corresponding to the ASTM 13-1160 5% point of the untreated feed. Thus, the isocraclcer product fraction boiling between 360 and 525 F. would have an undesirably high initial freezing point (-30 F), whereas the portion of the isocracker effluent boiling from 360 to the temperature represented by the ASTM D-l160 5% of the initial feed had an initial freezing point of less'than -70 F, well below the desired -40 F.

8 Example 3 Feedstock B was reduced in nitrogen content by a multistage hydrofining operation involving contacting a portion of the raw feedstock in a first hydrofining zone, along with 4500 s.c.f. of hydrogen per barrel of feed, with the same hydrofining catalyst as employed in Example 1 at a pressure of 1200 p.s.i.g., a maximum catalyst temperature of 640 F. and an 0.4 LHSV. The remaining portion of untreated feedstock B was then contacted, along with 6000 s.c.r". of hydrogen per barrel of feed, with a hydrofining catalyst composed of 6.5 weight percent (as the metal) molybdenum oxide and 2.7 weight percent (as the metal) cobalt oxide disposed on an alumina support at a pressure of 1200 p.s.i.g., a maximum catalyst temperature of 730 F, and an 0.4 LHSV. The two separately hydrofined portions were then blended and repassed through the first hydrofining stage, along with 4100 s.c.f. of hydrogen per barrel of feed, over the same first stage catalyst at a pressure of 1200 p.s.i.g., a maximum catalyst temperature of 675 F. and an LHSV of 0.6. The properties of the denitrified, normally liquid efiluent from the multistage hydrofining operation described above are given in Table IV below.

TABLE IV Gravity, API 31.2 Aniline point, F. 180 Distillation ASTM D-1160, F:

Start/5% 427/536 10% 572 30% 643 50% 690 70% 730 780 %/end point 801/845 Sulfur, ppm 8 Total nitrogen, p.p.m 1.3 Pour point, F +60 The normally liquid efiluent from the hydrofining operation, along with a recycle stream boiling above 525 'F., were contacted in an isocracking zone in exactly the same manner as described in Example 1 except that the isocracking reaction temperature was in the range of from 568 to 578 F., the on-stream period was hours, the per-pass conversion was 60 volume percent to products boiling below 525 F. and the 525 F. plus portion of the isocracker eflluent was recycled thereto. Hydrogen consumption was in excess of 500 s.c.f. per barrel of isocracker feed converted to products boiling below 525 F. The yields of the isocracker effluent boiling below 525 F. are given in Table V below.

TABLE V Ylelds Based on Fresh Feed Product Weight Percent Volume Percent resented by the ASTM 13-1160 point of the untreated feed, namely, 631 F, has an initial freezing point well above '30 F. Thus, it is apparent that with this particular feedstock, the desired freezing point fuel must have an end point of 525 F. rather than the temperature represented by the 5% distillation point (631 F.) of the initial feed.

Feedstock C was pretreated by contact, along with 6000 s.c.'f. of hydrogen per barrel of feed, with the hydrofining catalyst of Example 1 at a pressure of 1200 p.s.i.g., a maximum catalyst temperature of 710 F., and an LHS-V of 0.5. The properties of the normally liquid efiluent from the single stage hydrofining zone are shown below in Table VI.

TABLE VI Gravity, API Aniline point, F Distillation ASTM D-1160, F:

Start/5% 357/541 Cir 95%/end point 763/799 Sulfur, ppm Total nitrogen, p.p.m Pour point, F

The normally liquid efiluent from the hydrofining, along with a recycle stream boiling above 525 lF., were contacted in an isocracking zone in an identical manner as described in Example 1 except that the reaction tempe-rature was 575 F, the on-strearn period was 173 hours, the hydrogen consumption was 857 s.c.f. per barrel of h-ydrofiner eflluent converted to product boiling below 525 F, the per-pass conversion was 60 volume percent to products boiling below 525 F, and the 525 F. plus portion of the isocracker effluent was recycled thereto. The yields of the hydrocarbon eflluent boiling below 5 F. are given in Table VII below.

TABLE VII Yields Based on Fresh Feed Product Weight Percent Volume Percent The portion of the isocracking zone eflluent boiling over the entire range of from 360 to 525 F, amounting to 32.3 volume percent of the fresh feed converted, 30.6 volume percent of the 0 synthetic portion of the efliuent and over 20 volume percent of the total effluent from the isocracker, had an initial freezing point of 7l P. The portion of the isooracking zone eflluent boiling from 360 F. to the temperature represented by the ASTM D-1l60 5% point of the untreated feed, namely, 643 F, has an initial freezing point considerably above 30 F. From these data it can be seen that, in order to produce a fuel boiling above 360 F. and having an initial freezing point below E, the portion boiling from 360 to 525 F. satisfies the freezing point specification, whereas the portion boiling from 360 F. to the temperature corresponding to the ASTM D1160 5% point of the initial feed (643 F.) has an initial freezing point above 40 F. Thus, it is apparent 10 that, with this. particular feedstock, the desired freezing point fuel must have an end point of 525 rather than the temperature represented by the ASTM =D-l 5% point of the untreated feedstock.

Example 5 TABLE VIII Yields Based on Fresh (Total) Feed Weight percent Product Volume percent LO c HMOCO C5 to F- 180 to 360 F 360 to 525 F The portion of the isocracking zone efiluent boiling from 360 to 525 F., amounting to 22.4 volume percent of the total feed to the isocracker and 34 volume percent of the C5+ synthetic portion of the effluent, had an initial freezing point of below -70 F. The portion of the isocracking zone effluent boiling from 360 F. to the ASTM 13-1160 5% point of the original untreated feed (643 F.) has a freezing point well above 30- F. Again, with this feedstock it can be seen that, when employing once-through operations in the isocracking zone, of the two different fuels represented by the boiling ranges 360 to 525 F. and 360 to 643 F. (the 5% distillation point of the untreated feed) only the former will have an initial freezing point of below 40 F.

The actual work-up of the effluent from the isocracking zone is largely dependent upon the fuel needs of the refiner employing the process. However, in general it is preferred to first pass the isocracker efiiuent into a high pressure separator operating under such conditions of temperature and pressure that a gaseous overhead is removed that :is essentially hydrogen. This can be recycled to the isocracking zone and/or the hydrofining zone if the latter is employed as the pretreating zone. The remainder of the efiluent can be stripped of any light gases present and is then preferably passed into one or more distillation zones wherein additional product separation is performed. Thus, a C fraction, containing a high proportion of isobutanes to normal butanes, can be recovered and passed to an alkylation zone or the like. Preferably a C to 180 F. fraction, likewise having an iso to normal paraffin ratio considerably above the equilibrium ratio, is recovered. This fraction is an excellent gasoline blending stock inasmuch as it will have an F-l leaded octane number of about 100. If desired,

'of course, the butanes can be included in this light, high since the inclusion, within the jet fuel, of the 325 to 360 F. portion of the isocracldng zone eflluent would not increase the freezing point but would, in virtually all cases, actually lower it still further. If a jet blending stock is required, then the intermediate fraction could boil in the range of from, say, 180 to 360 F. In any case, the intermediate fraction recovered from the efiiuent of the isocracking zone is a superior reforming feedstock in that it contains a high proportion of cyclic compounds that will, upon reforming, produce an excellent, high octane gasoline. It might also be noted that the end point of the intermediate fraction can be raised to 400 F., or thereabouts, and still be a superior reformer feed.

As described above, a fuel fraction, boiling above the intermediate fraction, can be recovered from the distillation zone. As has been shown in the examples, this fuel cut is characterized by a low initial freezing point. The boiling range of this fuel fraction can be varied to meet the refinery and/or seasonal needs of the refiner. Thus, the fuel fraction can be one or more fuels, such as jet blending stocks, full range jet fuels, diesel fuels,

and higher boiling heating oils and the like. While these fuels, as actually separated, may have properties which differ from those of the jet fuels and jet blending stocks which are the prime object of this invention, that fraction boiling within the range within the scope of this invention will have an initial freezing point below about ---40 F.

A bottoms fraction can also be recovered from the isocracking zone efiluent, the actual boiling range being a function of the feed boiling range and the breadth of the lighter fuel fraction boiling range. All, or portions, of any bottoms fraction can be recycled to the isocracking zone, used as heavy fuels or as a superior feedstock for a catalytic cracking unit.

It is obvious that many modifications can be made in the operation and recovery techniques and still be within the scope of the present process.

We claim:

In a hydrocrack-ing process wherein a hydrocarbon distillate and at least 1500 s,c.f. of hydrogen per barrel of said distillate are contacted in a hydrocracking zone, with a net hydrogen consumption and at a per-pass conversion to products boiling below the initial boiling point of said distillate of at least volume percent, with a hydrocracking catalyst selected from the group consisting of compounds of nickel and cobalt intimately associated with an active, solid, acidic catalyst support at a hydrogen partial pressure of at least 350 p.s.i.a., at hydrocracking temperatures, and at a liquid hourly space velocity above 0.2 v./v./hr., the method of producing with said process a fuel of superior quality boiling generally above the gasoline range, which comprises: (a) selecting as the hydrocarbon distillate feed to said process a stock containing nitrogen, boiling over a range of at least F., and having a 5% boiling point above about 475 F.; (b) hydrofining said feed stock as necessary to reduce the nitrogen content thereof to below about 10 ppm. total nitrogen; (c) initiating said contacting at a temperature from 350 to 600 F, (d) raising the contacting temperature as necessary during the on-stream period to maintain said per-pass conversion; (e) withdrawing an effluent from said hydrocracking zone; (1) recovering from said effluent as a superior quality fuel useful for jet or diesel fuel purposes a fraction having an end point below 525 F. and boiling between about 360 F. and the 5% point of the feed, said fraction being characterized by an initial freezing point of less than about 4-0 F. and being essentially free of normal paraffins.

References fired in the file of this patent UNITED STATES PATENTS 2,428,692 Voorhies Oct. 7, 1947 2,944,006 Scott July 5, 1960 2,952,626 Kelly et al Sept. 13, 1960 

